Design and control of a hybrid reactive extraction configuration for the production of tert-butyl alcohol

Journal of Cleaner Production 239 (2019) 118018

Contents lists available at ScienceDirect

Journal of Cleaner Production journal homepage: www.elsevier.com/locate/jclepro

Design and control of a hybrid reactive extraction configuration for the production of tert-butyl alcohol Felicia Januarlia Novita, Bing-You Zou, Hao-Yeh Lee* Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei, 10607, Taiwan

a r t i c l e i n f o

a b s t r a c t

Article history: Received 16 November 2018 Received in revised form 8 August 2019 Accepted 11 August 2019 Available online 13 August 2019

In this study, the design and control of the hydration process for the production of tert-butyl alcohol (TBA) through a hybrid reactive extraction configuration was proposed. Firstly, TBA production by using a reactive distillation configuration and a hybrid reactive extraction configuration without the inert nbutene (NC4) was investigated. Due to the azeotrope between TBA and water (H2O), azeotropic distillation is required in the separation section with cyclohexane (CyH) as the entrainer. Compared to the reactive distillation configuration, the hybrid reactive extraction configuration was superior in terms of the total annual cost (TAC) and energy savings for the production of 99.9 mol% of TBA. Subsequently, the feed composition with the inert NC4 for a more realistic feeding condition was also investigated while using the hybrid reactive extraction configuration. The major differences in the two feed conditions were the compositions in the outlet of the reactive extraction section and the functions of the C1 column. Then, two control structures [a single-point temperature control (CS1) and a dual-point temperature control (CS2)] with three inventory control loops (inventory AeC) of the hybrid reactive extraction configuration with the inert NC4 were proposed. The simulation results revealed that both stoichiometric balance and liquideliquid equilibrium (inventory C) in the RE column should be used for the dynamic operation. The CS2 with inventory C (CS2C) was found to be the optimal control strategy in this process. © 2019 Elsevier Ltd. All rights reserved.

Handling editor. Jin-Kuk Kim Keywords: Reactive distillation Reactive extraction Tert-butyl alcohol Dynamic simulation Energy efficiency

1. Introduction Process intensification is a chemical and process design method that leads to more effective and efficient processes in terms of energy and installation costs. Process intensification has been gaining considerable attention across chemical and process industries, especially in the petroleum refining industry, which is the largest industrial consumer of energy (U.S. Energy Information Administration, 2018). The global oil demand has been increasing annually, as presented in Fig. 1 (International Energy Agency, 2018). This situation has led to the demand for innovation to minimize energy consumption of the process. Reactive distillation, divided wall column, and heat integrated distillation column are several excellent applications of process intensification technologies that have been extensively studied to date (Kiss and Jobson, 2018). Reactive distillation is a well-known technology that combines two processes (reaction and separation) into a single column equipped with catalytic packing. Equilibrium-limited reactions

* Corresponding author. E-mail address: [email protected] (H.-Y. Lee). https://doi.org/10.1016/j.jclepro.2019.118018 0959-6526/© 2019 Elsevier Ltd. All rights reserved.

provide higher yields by separating the reactants from the products. Moreover, the heat of the reaction is applied directly to the separation process. The advantages of using reactive distillation are as follows: requirement of fewer equipment, lower capital cost, lower energy use, and higher product yields, especially for equilibrium-limited reactions (Novita et al., 2015). These advantages of reactive distillation are closely associated to the principles and challenges of green engineering (Malone et al., 2003). The principles of green engineering have been recently described in the context of design, both for manufacturing molecules in chemical processes and for more general products and systems (Malone et al., 2003). One of the green engineering principles is innovation in processes, procedures, and materials used. As aforementioned, reactive distillation can be a green engineering technology for the vaporeliquid separation process with chemical reaction. However, for other combinations of chemical transformations with separations such as for liquideliquid separation, the extraction process can be considered. The latest trends in extraction techniques have largely focused on finding solutions that minimize the use of solvents. It seems like the extraction techniques are currently moving in the direction of green engineering

2

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Nomenclature ATV AQ C1 C2 C4 CyH CS1 (A)

auto-tune variation aqueous phase azeotropic distillation column stripping column butene cyclohexane single-point temperature control in C1 column (with inventory A) CS2 (A-C) dual-point temperature control in C1 column (with inventory A-C) FC flow controller H2O water i-C4 iso-butene k second-order kinetic constant (m6/kmol/kg/s) Kc chemical equilibrium constant (m3/kmol) KA inhibition constant for TBA (m3/kmol) LC level controller LLE liquid-liquid equilibrium n payback period N.B.P normal boiling point (
C)

principles. Later, design and innovation in the extraction processes would lead to reduction in energy consumption, allow the use of several alternative solvents and renewable natural products, and guarantee a safe and high-quality extract or product (Chemat et al., 2012). Processes that combine reaction and separation systems have been extensively studied for reactive distillation (Lee et al., 2012; Novita et al., 2018a, 2018b; 2017; Wang et al., 2008). Other researchers have also proposed several alternative intensified processes to enhance the energy and cost savings such as combining reactive distillation with thermally coupled arrangement (Ojeda et al., 2012; Novita et al., 2015) and combining distillation with pervaporation (Norkobilov et al., 2017). However, reactive extraction, which is also a chemical process that involves reaction and separation processes that occur simultaneously in a single processing unit, are rarely discussed. The primary difference between reactive extraction and reactive distillation is the type of separation involved. For reactive extraction, the liquideliquid phase or solideliquid phase separation is involved. Imam et al. (2013) proposed a series of reactors and decanters for indirect hydration of cyclohexene. This process operation was theoretically a reactive

NC4 n-butene NF C1 feed stage location Nrxn number of reactive stages NT number of total stages OR organic phase P proportional PC pressure controller PI proportional-integral Q reboiler duty (kW) RD reactive distillation column RE reactive extraction column RR reflux ratio TAC total annual cost (US$) TACoperating operating costs (US$) TACcapital equipment costs (US$) TBA tert-butyl alcohol TC temperature controller UNIQUAC universal quasichemical VLE vapor-liquid equilibrium VLLE vapor-liquid-liquid equilibrium X mole fraction DT temperature difference (
C)

extraction process. Furthermore, Garcia Jurado et al. (2013) used a similar concept to propose biodiesel production using a hybrid reactive extraction configuration. Their study reported the following advantages of reactive extraction: high conversion due to excess of reactants at both ends, low temperature operation, and high energy saving (approximately 83%) by using liquideliquid extraction. In this study, the application of reactive distillation and hybrid reactive extraction for the hydration of the iso-butene (i-C4) and water (H2O) process to produce tert-butyl alcohol (TBA) was investigated and compared. The source of i-C4 is usually raffinate C4, which is a byproduct of the extractive distillation process of 1,3butadiene. In other words, TBA production becomes a major route for isolating i-C4 from raffinate C4. Moreover, TBA can be used to improve antiknock and antifreeze characteristics when added to gasoline (Sada et al., 1981). TBA is also a crucial chemical intermediate used to produce methyl tert-butyl ether and ethyl tertbutyl ether. Based on the uses of TBA, it can be evidenced that TBA is an important component in the chemical industry. The implementation of reactive distillation and hybrid reactive extraction configurations into the TBA production process is one of the primary contributions of this study. Two different feed conditionsdwithout and with the inert n-butene (NC4), were applied when the hybrid reactive extraction configuration was used. The inert NC4, therefore, will not react with H2O, and it is also desirable to know if the hybrid reactive extraction configuration with the inert NC4 could be controlled properly for practical use. Thus, the overall control strategy of the hybrid reactive extraction configuration with the inert NC4 will be another main contribution of this study. Commercial simulatorsdAspen Plus V8.6 and Aspen Plus Dynamics V8.6dwere used for steady-state and dynamic simulations, respectively. 2. Kinetic model and thermodynamic data 2.1. Kinetic model

Fig. 1. Daily demand for oil worldwide from 2005 to 2018 (million barrels per day).

The liquid hydration reaction of i-C4 and H2O is processed as follows:

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018 Table 1 Kinetic model and parameters for the studied system. Kinetic model

Parameter [unit]

C kðCieC4 CH2 O  TBA Þ KC R ¼ 1 þ KA KC

k ¼ Exp (15.03e8844/T) [m6/kmol/kg/s] Kc ¼ Exp (3160/T-6.78) [m3/kmol] KA ¼ Exp (26.6e8540/T) [m3/kmol]

This exothermic reversible reaction was catalyzed using the heterogeneous catalyst Amberlyst® 15. The kinetic model of the aforementioned reaction was found in the study by Velo et al. (1988). The rate of the liquid hydration reaction is usually low due to the immiscibility of reactants. A high concentration of TBA leads to a reverse reaction. However, the high concentration can improve the solubility of i-C4 and H2O, thus increasing the direct reaction rate (Caceres et al., 1988; Dellon et al., 1988; Zhang et al., 2003). Information of the kinetic model and parameters in Eq. (1) are presented in Table 1. In Table 1, k is the second-order kinetic constant, Kc is the chemical equilibrium constant, and KA is the inhibition constant for TBA. Note that the unit of the kinetic constant (k) is m6/kmol/kg/s. This unit should be converted to m3/kmol/s in the simulation study by assuming a catalyst density of 770 kg/m3 and by assuming that the catalyst occupies one-half of the liquid holdup in the reactive tray. In this study, a Fortran subroutine was used for the reaction kinetic. 2.2. Thermodynamic data In this study, two types of reactive systems were used. The first one was a reactive system for the reactive distillation and hybrid reactive extraction without the inert NC4, and the second one was a

3

reactive system for the hybrid reactive extraction configuration with the inert NC4. Table 2 shows all components involved in each process. In the system without the inert NC4, there were a total of four azeotropes; one each between H2O and TBA; between H2O and CyH; between TBA and CyH; and between H2O, TBA, and CyH. While there were a total of five azeotropes found in the system with the inert NC4, of which four azeotropes were the same as those in the system without the inert NC4. The additional azeotrope was between H2O and NC4. In this study, the universal quasichemical (UNIQUAC) activity coefficient model was used for the vaporeliquid equilibrium (VLE) and the vaporeliquideliquid equilibrium (VLLE), and the Aspen Plus built-in binary parameters by Aspen Technology are listed in the Supplementary Data. The boiling point of the pure components and azeotropes observed in this study are presented in Table 3. In this study, extraction and azeotropic distillation were conducted in the process of both reactive systems, thus the liquideliquid equilibrium (LLE) was crucial. The liquideliquid diagrams of i-C4/H2O/TBA, NC4/H2O/TBA, and H2O/TBA/CyH are presented in Fig. 2. Fig. 2a reveals that i-C4/H2O/TBA has a suitable two-phase zone in the three-phase diagram; thus, the reactive extraction process can be designed. By confirming the azeotropic point and verifying thermodynamics, it was found that H2OeTBA had a co-boiling point. In Yu et al.‘s study (Yu et al., 2015), CyH was added to azeotropic distillation as an entrainer. Fig. 2b reveals that NC4 and i-C4 phase diagram results were very similar. Fig. 2c presents a three-phase diagram of H2O/TBA/CyH at a temperature of 60
C and a pressure of 1 bar. The azeotropic point of the three components was located in the two-phase region. Therefore, the top portion of the distillation column must be brought into the phase separation tank to achieve separation. 3. Conceptual design of the TBA production process In the study by Garcia Jurado et al. (2013), the hybrid reactive extraction design was applied to the transesterification of triglycerides for biodiesel production. The energy saving effect was quite good, and the design of the hybrid reactive extraction configuration was quite promising. The hydration reaction of i-C4 and H2O is good for implementation in hybrid reactive extraction

Table 2 Components involved in both systems e without and with the inert processes. System

Process

Component Reactant

Product

Entrainer

Inert

Without the inert NC4

Reactive distillation Reactive extraction Reactive extraction

i-butene (i-C4) and water (H2O)

tert-butyl alcohol (TBA)

cyclohexane (CyH)

e

i-butene (i-C4) and water (H2O)

tert-butyl alcohol (TBA)

cyclohexane (CyH)

n-butene (NC4)

With the inert NC4

Table 3 Normal boiling point and azeotropes observed in this study. Normal boiling point (N.B.P) Component NC4 i-C4 CyH TBA H2O a b

Pressure at 435.75 kPa. Pressure at 101.33 kPa.

Azeotropes


N.B.P ( C)

Components

Comp.

Temp. (
C)

6.24 6.90 80.72 82.42 100.00

a

1.44/98.56 24.5/20.7/54.8 30.0/70.0 36.2/63.8 37.1/62.9

36.99 65.31 69.41 71.89 80.30

H2O/NC4 H2O/TBA/CyH b H2O/CyH b TBA/CyH b H2O/TBA b

4

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 2. LLE diagram for (a) i-C4/H2O/TBA (60
C, 10 bar); (b) NC4/H2O/TBA (60
C, 10 bar); (c) H2O/TBA/CyH (60
C, 1 bar).

process because it is a low temperature reaction with suitable thermodynamic properties. Therefore, in this study, the TBA production process was selected for conducting a case study in which the hybrid reactive extraction configuration was implemented. This case was compared with a case including the reactive distillation configuration. By considering the presence of the inert in the raw materials and assuming that the inert was NC4, the hybrid reactive

extraction configuration with the inert NC4 was also investigated. 3.1. TBA production process without the inert NC4 3.1.1. Reactive distillation configuration Fig. 3 presents a detailed process flowsheet of the reactive distillation configuration for the TBA production process without

Fig. 3. Flowsheet and results of the reactive distillation configuration for the TBA production process without the inert NC4.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

the inert NC4. For the reactive distillation configuration, the process was divided into two sections. The first section was the reactive section comprising an RD column, and the second section was the separation section comprising a C1 azeotropic distillation column and a C2 stripping column. In the design of the RD column, a total reflux was proposed to avoid the drawing out of the light reactant i-C4 from the top of the column. The entire upper section of this RD column was the reactive section. The i-C4 and H2O feeds, including fresh H2O and recycle streams, entered into the reflux drum of the RD column. A combination of fresh H2O and recycle streams increased the conversion of i-C4 to TBA. However, the bottom of the RD column contained the TBA/H2O mixture due to the excess H2O that was recycled back into the RD column. This implies that the excess H2O was the most sensitive variable. Due to the azeotrope between TBA and H2O, a change in this design variable causes a relatively large change in the reboiler duty of the RD column. Thus, an optimal excess H2O flowrate should be found. Subsequently, the bottom product of the RD column, which was TBA and H2O, entered the separation section. An additional C1 heterogeneous azeotropic distillation with CyH as the entrainer was required due to the azeotrope between TBA and H2O. Moreover, a C2 stripping column was required at the end of the separation section for recycling excess H2O into the reactive section. This action helped the reaction to progress in the positive direction under the excess H2O condition. The makeup flow of CyH was added to the top decanter to maintain the material balance of the entire process. Due to thermodynamics constraints, i-C4 concentration in the bottom of the RD column was lower than 1 ppm in order to prevent the accumulation of i-C4 in the top decanter of the C1 heterogeneous azeotropic distillation column. In other words, i-C4 was a limited reactant and should completely react in the RD column. The primary product of this process was 99.9 mol% of TBA, which was obtained in the bottom stream of the C1 column. The total annual cost (TAC) analysis was performed to determine the optimal design of the TBA production process. Eq. (2) expresses the TAC, which consists of the operating costs (TACoperating) and equipment costs (TACcapital), at a payback period (n) of 3 years.

TAC ¼ TACoperating þ

TACcapital 3

(2)

The catalyst cost was included in the operating costs. The life of the catalyst was assumed to change every 3 months. The price of Amberlyst® 15 was approximately US$3.5 per pound. TAC calculated on the basis of a study by Douglas (1998) by ignoring the costs of piping and pumps for simplicity, as detailed in the Supplementary Data. For the prices of the low-pressure steam and cooling water, referred to Turton et al. (2009). Six variables were optimized using an iterative optimization procedure: (1) the number of reactive stages in the RD column (Nrxn,RD), (2) number of total stages in the RD column (NT,RD), (3) iC4 feed location, (4) excess H2O flowrate, (5) number of total stages in C1 (NT,C1), and (6) number of total stages in C2 (NT,C2). Fig. 4 presents the sensitivity plot of those variables. Even though an iterative optimization procedure does not warranty the global optimal solution and others optimization methods can do a global optimization strategy, such as shortest separation lines (Lucia et al., 2006). However, the iterative optimization procedure was still chosen in this study because this procedure gives benefits as it can be implemented to other processes and the engineers can investigate the sensitivity of the design variables and the production

5

phenomena on the basis of their knowledge (Luyben and Yu, 2008; Chung et al., 2015; Novita et al., 2017). As aforementioned, all feedsdi-C4 and H2Odwere fed into the reflux drum of the RD column. The reactant H2O was fed into the reflux drum because it has the highest boiling component in the system. The reactant i-C4 was the lightest component in the system and was fed into the lower half of the RD column (Dellon et al., 1988). However, the feed location of i-C4, which is the reflux drum (feed position ¼ 0), contributed minimum to TAC (Fig. 4). Fig. 4 reveals that when the number of total stages in the RD column was 59, i-C4 concentration in the bottom outlet of the RD column was approximately 0.9 ppm. Later, the number of total stages in the RD column was set to 58 to meet the constraints of the i-C4 concentration in the bottom outlet of the RD column. The reactive stages were from the top to the 52nd stage, which was determined by the temperature limit of the catalyst. Fig. 5a and b presents the relation between excess H2O and the duty of the RD column and present the T-xy diagram of H2OeTBA, respectively. The energy requirement of the RD column decreased with a decrease in the flowrate of H2O. However, the energy consumption increased sharply when the flowrate of H2O was set to 55 kmol/h. When the operating pressure of the RD column was 6 bar and the excess H2O flowrates were 55 kmol/h and 65 kmol/h, the H2O concentration in the bottom outlet of the RD column was approximately 35.5 mol% and 39.4 mol%, respectively. The concentration gradually shifted toward to the two-phase area, thus causing the reboiler of the RD column to consume a high amount of energy. Therefore, the excess H2O flowrate was set at 65 kmol/h as the optimum flowrate. In the separation section, the C1 azeotropic distillation column and the C2 stripping column were optimized independently for conducting the analysis. In the C1 azeotropic distillation column, as the reflux was very small, the C1 feed stage location (NF,C1) was not considered. However, the number of total stages in C1 (NT,C1) was optimized. Fig. 4 reveals that the TAC decreases as NT,C1 decreases. The optimum NT,C1 value was 37 stages, and the C1 feed location was on the 11th stage. The optimum number of total stages in C2 (NT,C2) was five. The number of total stages used was less because more energy was required as an effect of the increase in the pressure drop of the stripper with a high number of total stages. Another reason can be that the H2O concentration in the feed stream of the C2 stripping column was already quite high. Moreover, as H2O was fed at the first stage, the number of total stages required in the C2 stripping column was less. The recovery rate of TBA in the C2 stripping column can be higher than 99.95% when only five stages are used; this is another reason for selecting NT,C2. In the RD column, as the main component in the condenser was the vapor of i-C4, it was crucial to increase the operating pressure of the column so that the condenser could use the less expensive cooling water. Fig. 6 presents the change in the boiling point of i-C4 under different pressures. Fig. 6 reveals that the boiling point of iC4 is approximately 48
C at a pressure of 6 bar. For setting an appropriate cooling water temperature, the operating pressure in the RD column was set at 6 bar. In the separation section, the operating pressure was set at 1 bar. Four assumptions were made to design the RD column: (1) the liquid holdup was set as half-full of the catalyst, (2) the maximum kinetic holdup was considered to be 10 times the tray holdup in the reflux drum, (3) the stage efficiency was 100% and the tray spacing was 0.6096 m, and (4) the feed pressure was 6 bar with a feed temperature of 35
C. The liquid composition and temperature profiles of the RD

6

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 4. Sensitivity plots of the design variables.

Fig. 5. (a) Analysis of the excess H2O flowrate for the RD column and (b) T-xy diagram of H2OeTBA at 6 bar.

column are presented in Fig. 7. The reactive zone temperature was lower than 120
C, and this situation obeyed to the temperature limit of the catalyst. The composition profile (Fig. 7b) reveals that two liquid phases occur in the reactive zone of the RD column. 3.1.2. Hybrid reactive extraction configuration The flowsheet of the hybrid reactive extraction configuration without the inert NC4 is presented in Fig. 8. In this system, the C4

feed only contained the reactant i-C4, and its flowrate was 100 kmol/h. By applying a concept similar to that proposed by Garcia Jurado et al. (2013), both reactantsdi-C4 and fresh H2Odwere countercurrent fed into the RE column. Unlike the process proposed by Garcia Jurado et al. (2013), H2O acted as a reactant and an extractive solvent in this hybrid reactive extraction design of the TBA production process. The operating conditions of the RE column were 10 bar and 60
C, which was set as per the lowest boiling point of i-C4 to avoid the appearance of vapor phase

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

7

Fig. 6. Change in the boiling point of i-C4 under different pressures. Fig. 9. Analysis of the total stages of the RE column and i-C4 conversion (%).

in the RE column. As the RE column simulated using series reactors and decanters, the column diameter was equal to the reactor diameter. The reactor volume was set at 2 m3. As a result, the column diameter was 2.05 m, and the height in each RE stage was 0.6096 m. Note that 99.5% conversion of i-C4 was obtained using six stages in the RE column, as shown in Fig. 9. Because the reaction was exothermic, no heat was supplied to the RE column. The solid line presented in Fig. 10a represents the liquid

composition of the organic phase (OR) and the dashed line represents the liquid composition of the aqueous phase (AQ). The composition of the H2O phase in each stage was maintained at a high value due to the tie line direction. If a higher amount of H2O is recycled, then the conversion of i-C4 and the liquid-liquid behavior will improve. Fig. 10b reveals the reactive ratio in each stage of the RE column. The reaction mostly occurred in the first stage of the RE

Fig. 7. (a) Temperature and (b) liquid composition profiles of the RD column.

Fig. 8. Flowsheet of the hybrid reactive extraction configuration without the inert NC4 for the TBA production process.

8

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 10. (a) Two liquid-phase composition profiles and (b) reactive ratio of the RE column in the hybrid reactive extraction system without the inert NC4.

Fig. 11. Comparison between hybrid reactive extraction and reactive distillation configurations without the inert NC4 in terms of (a) energy requirement and (b) TAC.

column. As in the reactive distillation configuration, two sections were involved in the hybrid reactive extraction configurationdthe reaction and separation sections. The RE column was the reactive section for the hybrid reactive extraction configuration. The separation section had two columns that operated at 5.5 and 1 bar. One column was a C1 stripping column and the other was a C2 heterogeneous azeotropic distillation column with CyH as the

entrainer. Both columns had to be maintained in a series arrangement due to the LLE and the unreacted i-C4 reactant in the downstream of the RE column. The unreacted i-C4 was recycled back to the RE column from the C1 stripping column. i-C4 concentration in the bottom stream of the C1 stripping column was controlled at lower than 1 ppm to prevent the accumulation of i-C4 in the downstream of the separation unit. After recycling the unreacted i-C4, the bottom stream of the C1 stripping column

Fig. 12. Flowsheet and results of the hybrid reactive extraction configuration with the inert NC4.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

9

Fig. 13. (a) Two liquid-phase composition profiles and (b) reactive ratio of the RE column in the system with the inert NC4.

Fig. 14. (a) Temperature and (b) liquid composition profiles of the C1 column for the hybrid reactive extraction configuration with the inert NC4.

Fig. 15. (a) Temperature and (b) liquid composition profiles of the C2 column for the hybrid reactive extraction configuration with the inert NC4.

consisted of the TBA/H2O mixture with a very low concentration of i-C4. Then, the mixture was fed into the top decanter of the C2 azeotropic distillation column. A heat exchanger unit was set at the downstream of the RE column and the bottom stream of the C1 stripping column to prevent wastage of energy. The OR outlet of the top decanter was recycled back to the C2 azeotropic distillation column, and the AQ outlet (H2O) was directly recycled to the RE column. Furthermore, the makeup flow of the entrainer was added to the top decanter to maintain the material balance of the entire process. Finally, 99.9 mol% of TBA was obtained from the bottom

stream of the C2 azeotropic distillation column. In the reactive distillation configuration, an iterative optimization procedure was used to optimize several design variables. The same principle was applied in the hybrid reactive extraction configuration. The optimum design conditions for the RE column were determined by analyzing the number of reactive stages. The same analyzing procedure was applied for the separation unit. The minimum TAC was used as the objective function of the optimization process.

10

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

contained more H2O than that from the reaction section of the reactive distillation configuration. Thus, an additional 14.3% and 24.5% of energy and TAC, respectively, was required in the separation section of the hybrid reactive extraction configuration compared to the reactive distillation configuration. Overall, the results revealed that the hybrid reactive extraction configuration without the inert NC4 could save up to 26.0% of the energy and 16.1% of the TAC compared to the reactive distillation configuration without the inert NC4. 3.2. TBA production process with the inert NC4

Fig. 16. Reaction tendencies of both systems: without and with the inert NC4.

3.1.3. Comparison of the reactive distillation and hybrid reactive extraction configurations without the inert NC4 After the analysis of both configurations, the design results are summarized in Fig. 11. The hybrid reactive extraction configuration could save 72.4% of the TAC in the reaction section and a high conversion of i-C4 could be achieved with no energy requirement due to the operation of the extraction process and the exothermic reaction. Due to the constraint of the LLE, the outlet stream from the reaction section of the hybrid reactive extraction configuration

3.2.1. Hybrid reactive extraction configuration The reactant i-C4 was mainly generated from the pyrolysis of light oil. However, in the light oil cracking process, complex low carbon alkanes were produced. After the separation process, mixed hydrocarbons were obtained, which mainly contained butane, butene, and butadiene. Butene was separated from butadiene through extractive distillation of 1,3 butadiene and mainly consists of 40e50 wt% of i-C4 (Blackbourn et al., 2010). By considering the presence of inert in the raw materials, assuming that the inert was 50 mol% of NC4, and considering the advantage of the hybrid reactive extraction configuration compared to the reactive distillation configuration, the hybrid reactive extraction configuration with the inert NC4 was investigated in this section. Fig. 12 presents the flowsheet of the hybrid reactive extraction configuration with the inert NC4. The mixed C4 feed contained 50 mol% of i-C4 and 50 mol% of NC4, and the total flowrate was 200 kmol/h. Both feedsdmixed C4 and fresh H2Odwere countercurrent fed into the RE column. The operating conditions of the RE column were 10 bar and 60
C, which

Fig. 17. Open-loop sensitivity test of the hybrid reactive extraction configuration with the inert NC4.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

11

Fig. 18. Basic control structuredCS1A.

was set according to the boiling points of i-C4 and NC4 to avoid the appearance of the vapor phase in the RE column. The column diameter was the same as in the process without the inert NC4, and there were six stages in the RE column. Fig. 13a presents the two liquid-phase composition profiles of the RE column. The solid line represents the liquid composition of OR, and the dashed line represents the liquid composition of AQ. Due to the inert NC4, the flowrates of the internal H2O and the

recycled H2O were higher in the system with the inert NC4 than in the system without the inert NC4. Fig. 13b presents the reactive ratio in each stage of the RE column. The reaction mostly occurred in the first stage of the RE column. Similar to the system without the inert NC4, there were two columns in the separation section for the system with the inert NC4dthe C1 column and the C2 heterogeneous azeotropic distillation column with CyH as the entrainer. Both columns were

Fig. 19. Dynamic responses in different temperature-controlled stages.

12

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

operated at 6 and 1 bar, respectively. The first column was for separating the C4 components, and the second one was for separating the product from the TBA/H2O mixture. The downstream of the RE column contained the C4 componentsdTBA and H2O. The C4 components could not react with i-C4 and the inert NC4. It was highly difficult to separate a small amount of i-C4 from NC4 because the boiling points of both components are very close to each other. Therefore, the unreacted i-C4 could not be recycled to the RE column and thus was directed to the C1 column

as a distillate together with the inert NC4. The yield of the product decreased because the unreacted i-C4 could not be recycled. However, because of an azeotrope between NC4 and H2O, the distillate included a small amount of H2O. A decanter was placed after the distillate to separate H2O from the C4 components. The C4 components were obtained at the OR outlet and H2O was obtained at the AQ outlet. The bottom stream of the C1 column was the TBA/ H2O mixture and was fed into the first stage of the C2 heterogeneous azeotropic distillation column. A heat exchanger was placed

Fig. 20. Dynamic response in CS1A: (a) throughput and (b) i-C4 composition disturbance.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

at the downstream of the RE column and the bottom stream of the C1 column to prevent energy wastage. The OR outlet of the decanter in the C2 azeotropic distillation column was recycled back to the C2 column, and the AQ outlet of the decanter was directly recycled to the RE column. Furthermore, the makeup flow of the entrainer was added to the top decanter of the C2 azeotropic distillation column to maintain the material balance of the entire process. Finally, 99.9 mol% of TBA was obtained from the bottom stream of the C2 azeotropic distillation column. Same as the previous configuration, an iterative optimization procedure was used to optimize several design variables in the reaction and separation sections. The minimum TAC function was used as the objective function of the optimization process. The optimized design conditions of the RE column were determined by analyzing the number of total stages of the RE column. In the system with the inert NC4, the number of total stages of the azeotropic column was the same as that in the system without the inert NC4. The detailed optimization results are presented in the Supplementary Data. The temperature and liquid composition profiles of the C1 and C2 columns for the hybrid reactive extraction configuration with the inert NC4 are presented in Figs. 14 and 15, respectively. In the liquid composition of the C1 column, the two liquid-phase zones were observed from the first stage to the seventh stage. 3.2.2. Differences between hybrid reactive extraction configurations without and with the inert NC4 The major differences between the two hybrid reactive extraction configurations (without and with the inert NC4) are the compositions in the outlet of the RE column and the function of the C1 column. The inert NC4 can cause a higher amount of recycled H2O to flow in the RE column so that the reaction conversion can improve. The reaction tendencies of both system studies are presented in Fig. 16. The points in the ternary diagram represents the composition of bulk liquid in each stage for the RE column. In the system without the inert NC4, the composition of each stage was located in the two liquid-phase regions, and the i-C4 concentration was very low in the final stage. The i-C4 concentration in the final stage was different for the system with the inert NC4. The concentration for the system with the inert NC4 was higher than that for the system without the inert NC4.

13

4. Dynamic simulation and control strategy development In this section, we discuss the control strategies for the hybrid reactive extraction configuration with inert NC4. The control strategies are meant to prevent the accumulation of the inert NC4 and unreacted i-C4 in the system and to maintain the purity of TBA at 99.9 mol%. There were two types of control loopsdinventory control and quality control. The inventory control was for the liquid level, flowrate, and pressure. This control was related to the material balance of the process and ensured process safety. The quality control was for temperature or composition and was used to maintain product purity. In this study, the product specifications were controlled using a temperature controller, which was used instead of a composition controller. The reason is that most composition analyzers, such as gas chromatographs, have large measurement delays and high capital and maintenance costs. 4.1. Design of the control loop In this study, the proportionaleintegral (PI) controller was used for flow, pressure, and temperature control loops. The proportional (P) controller was used for level control. The measurement points of the temperature in the two distillation columns were identified by conducting the open-loop temperature sensitivity test. The controlled stage temperature should have the largest gain toward the manipulated variables. Furthermore, the auto-tune variation (ATV) method was used to tune the controller parameters in Aspen Plus Dynamics. The ATV initiates from the relay-feedback method, which can enable us to obtain the ultimate gain and ultimate period. Then, the controller parameters were calculated using TyreusLuyben tuning relations. The sequential iterative tuning procedure was used to determine the final controller parameter. Moreover, three first-order time lags, which were set to 0.5 min, were added to each temperature control loop to improve the practicality of the simulation. All the detailed parameters of the temperature controller are presented in the Supplementary Data. To identify the stage with the highest temperature sensitivity toward manipulated variables, an open-loop temperature sensitivity test was conducted. A reflux ratio of ±0.01% and reboiler duty of ±0.005% were set for the C1 column, and a reboiler duty of ±0.1% was set for the C2 heterogeneous azeotropic distillation column.

Fig. 21. Improved control structuredCS2A.

14

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

When the steady state was reached, the temperature sensitivity was analyzed. The stages with the largest temperature variation were selected for temperature control because they are the most sensitive. The results are presented in Fig. 17. After the control structures were developed, they were tested with throughput and composition disturbances to observe their abilities of rejecting disturbances. Throughput disturbances of ±10% were added. Because feeds were assumed to be 50% of NC4 in the steady-state design, composition disturbances included ±5% iC4 composition change.

4.2. Basic control structuredCS1 with inventory A The basic control structure is presented in Fig. 18. In the RE column, two liquid levels have to be controlled. The first is the OR level of the RE column top, and it is controlled using the flowrate of the outlet stream of the RE column. The second is the AQ level of the RE column bottom, and it is controlled by the fresh water inlet to avoid the snow-ball effect in the RE column bottom. The RE column was simulated using a series reactors and decanters, and the column temperature was maintained under an isothermal

Fig. 22. Dynamic response in CS2A: (a) throughput and (b) i-C4 composition disturbance.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

15

Fig. 23. Improved control structuredCS2B.

condition by using a jacketed heat exchanger. However, it was not possible to simulate this type of column by using a jacketed heat exchanger in the simulator. Therefore, a temperature controller was used in each reactor to maintain the RE column under an isothermal condition. In C1 and C2 columns, the sump level was controlled using the bottom flowrate. In the C1 column, the top pressure was controlled using the condenser duty, and the reflux drum level was controlled using the distillate. In the C2 column, the top pressure was controlled by the valve between the column top and the condenser, and the two liquid levels in the top decanter of the C2 column were controlled by the flowrate of the organic and aqueous streams. The decanter temperature should be controlled by the condenser if the outlet temperature of the condenser changes. In the section of the quality control loop, the reflux ratio of the C1 column was fixed and the temperatures of the 12th stage in the C1 column and the 9th stage in the C2 column were controlled by

the reboiler duties of the C1 and C2 columns, respectively. As there was a single-point temperature control in the C1 column, the structure was termed CS1. However, to maintain a stoichiometric balance between the two reactants, the ratio of the flowrate of the total H2O inlet to the flowrate of outlet product TBA should be maintained; this inventory control loop was known as inventory A. Therefore, this basic control structure was termed CS1 with inventory A (CS1A). In the open-loop temperature sensitivity test, the result reveals that the largest temperature variation appears in the 7th stage and not in the 12th stage. Fig. 19 presents the dynamic responses under the þ5% composition disturbance for different C1-controlled stages. However, when the 7th stage was selected, poor dynamic responses were observed. Fig. 20 presents the dynamic responses under throughput disturbances and i-C4 composition disturbances. All types of disturbances were input into the process at the third hour. When 10%

Fig. 24. Improved control structureeCS2C.

16

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 25. Dynamic response of the throughput disturbance in CS2AeCS2C: (a) flowrate, (b) temperature, and (c) purity.

throughput disturbances were input for testing, the steady-state value deviation of the NC4 purity was 0.1%. When þ5% i-C4 composition disturbances were input, the steady-state value of the NC4 purity was lower than 98.0 mol% and increased the reboiler duty of the C1 column. On the other hand, there is a well response for þ10% throughput and 5% i-C4 composition disturbances. The final steady-state value of the TBA purity for all types of disturbances could be rejected even with a very small steady-state value deviation. 4.3. Improved control structure by using different temperature control loopsdCS2 with inventory A In CS1, there was only one temperature controller to maintain the purity of the outlet stream in the C1 column. Although the final TBA purity could be rejected even with very small deviation in all types of disturbances, it could cause a large deviation in the NC4 purity and wastage of TBA. Therefore, another control structure was proposed to overcome these limitations. In this new control structure, two temperature controllers (at 7th and 12th stages) in the C1 column were used due to the two peaks in the open-loop temperature sensitivity of the C1 column. The inventory control loop used was inventory A, same as that used in CS1. Due to this type of arrangement, the new control structure was termed CS2 with inventory A (CS2A). The temperatures of the 7th and 12th stages were controlled by the top reflux rate and the reboiler duty of the C1 column. The improved control structure is presented in Fig. 21. Fig. 22 presents the dynamic responses under throughput and iC4 composition disturbances. All types of disturbances were input into the process at the third hour, and the disadvantages of the basic structure could be identified using the improved structure.

The NC4 purity could be rejected. The final steady-state value of TBA purity for all types of disturbances could be rejected even with a very small steady-state value deviation. 4.4. Improved control structure by using different inventory control loops In the conventional chemical process, there is a stoichiometric balance in the reaction section. Thus, one of the crucial issues in the suitable control structure is to maintain the stoichiometric balance when the disturbances occur. However, in the reactive extraction process, the chemical reaction and extraction occur in a single unit. This implies that the stoichiometric balance and LLE should be maintained in the reaction section. To identify the best inventory control loop, three types of inventory control loops (inventory AeC) were proposed by different ratio controllers to maintain the stoichiometric balance, LLE, and both in this study. The quality control loop adopted CS2, and three types of inventory control loops were proposed to obtain different goals. The first one was the previous inventory useddinventory A, which used the ratio of the flowrate of the total H2O inlet to the flowrate of the outlet product TBA to maintain a stoichiometric balance between the two reactants. Second one was inventory B, which used the ratio of the flowrate of the total H2O inlet to the flowrate of the outlet C4 components to maintain the LLE in the RE column. The last one was inventory C, which used the ratio of the flowrate of the total H2O inlet to the flowrate of the outlet RE column to maintain both stoichiometric balance and LLE in the RE column. Figs. 23 and 24 present the control structure with inventories B and C. The results of the control structure for different inventory

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 25. (continued).

17

18

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 26. Dynamic response of the i-C4 composition disturbance in CS2AeCS2C: (a) flowrate, (b) temperature, and (c) purity.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

19

Fig. 26. (continued).

control loops with throughput and i-C4 composition disturbances are presented in Figs. 25 and 26. Fig. 25 reveals that all control structures had similar performance under the throughput disturbance. Fig. 26 reveals that the flowrate, temperatures, and compositions of the i-C4 composition disturbances differed. According to the flowrate results, the responses of the H2O inlet flow and the recycled H2O flow present that the performance of CS2C was slightly better than that of others. After the disturbances occurred, the process could reach a steady state again within 6 h. In the temperature results, the responses of TC1 and TC2 revealed that the dynamic response of CS2C was faster than that of others. In the composition result, the responses of the NC4 purity revealed that the performance of CS2C was slightly better than that of others. From the simulation, it can be summarized both stoichiometric balance and LLE should be maintained in the RE column. Fig. 27 presents the composition variation in each component at each stage when composition disturbance occurs. The figure explains why CS2C is the optimal control structure in this study. The variation in TBA, C4, and H2O purity represents stoichiometric balance, LLE, and overall balance, respectively. The variation in TBA purity in CS2A is the least in Fig. 27a, and the variation in C4 purity in CS2B is the least in Fig. 27b. Fig.. 27a and b reveal that the variation in the TBA and C4 purity in CS2C is in between CS2A and CS2B. The variation in the H2O purity in CS2C is the smallest in Fig. 27c. These results explain that the overall balance should be maintained in the RE column. 5. Conclusions In this study, the designs of the reactive distillation and hybrid reactive extraction configurations for TBA production were studied and compared. An additional step of azeotropic distillation with CyH as the entrainer was used for the separation process to

overcome the azeotrope between TBA and H2O and meet the specification of 99.9 mol% of TBA. Among both configurations, the hybrid reactive extraction configuration without the inert NC4 was superior to the reactive distillation configuration under the same feed condition. The hybrid reactive extraction configuration could save up to 26.0% of the energy and 16.1% of the TAC compared with the reactive distillation configuration. Then, the hybrid reactive extraction configuration with the inert NC4 was also investigated. The major differences between the processes without and with the inert NC4 in the hybrid reactive extraction configuration were the compositions in the outlet of the RE column and the function of the C1 column. Effective control structures were also proposed for the hybrid reactive extraction configuration with the inert NC4. The control structures were CS1 (a single-point temperature control in the C1 column) and CS2 (a dual-point temperature control in the C1 column). Because the chemical reaction and liquideliquid separation occurred in the RE column, three types of inventory control loops (inventory AeC) were studied to maintain a stoichiometric balance, LLE, and both. Variations in the feed composition and throughput changes were investigated. Product purities were maintained quite close to their specifications. From the dynamic results, it revealed that the optimal control structure is CS2C as it provided the fastest responses to reach the steady state under both disturbances. Overall, the proposed hybrid reactive extraction configuration can be a promising technology for TBA industrial application owing to the real-feed composition (with the inert NC4) applied in this study. Moreover, several energy integration configurations such as thermally coupled distillation or external heat-integration may be implemented with the hybrid reactive extraction configuration to enhance the energy and cost savings. These energy integration topics can be discussed in the future.

20

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018

Fig. 27. Variation in each composition in CS2AeCS2C: (a) TBA, (b) C4, and (c) H2O.

Notes The authors declare no competing financial interest. Acknowledgments This financial supports from the Ministry of Science and Technology of Taiwan under grant MOST 104-2221-E-011-149-MY2 and MOST 107-2221-E-011-062, which is gratefully acknowledged. Appendix A. Supplementary data Supplementary data to this article can be found online at

https://doi.org/10.1016/j.jclepro.2019.118018. References Aspen Technology, Aspen Plus Built-In Parameter, Version 8.6. Blackbourn, R.L., Bolinger, C.M., Potter, M.W., 2010. The Production of 1-hexene from Butadiene Raffinate-1. US Patent WO2010138435A1. Caceres, E., Puigjaner, L., Recasens, F., 1988. A trickle-bed process for hydration of isobutene to tert-butyl alcohol: a study of reactor performance. Chem. Eng. J. 37, 43e52. https://doi.org/10.1016/0300-9467(88)80005-4. Chemat, F., Vian, M.A., Cravotto, G., 2012. Green extraction of natural products: concept and principles. Int. J. Mol. Sci. 13, 8615e8627. https://doi.org/10.3390/ ijms13078615. Chung, Y.H., Peng, T.H., Lee, H.Y., Chen, C.L., Chien, I.L., 2015. Design and control of reactive distillation system for esterification of levulinic acid and n-butanol. Ind. Eng. Chem. Res. 54, 3341e3354. https://doi.org/10.1021/ie500660h.

F.J. Novita et al. / Journal of Cleaner Production 239 (2019) 118018 Dellon, A., Torck, B., Hellln, M., 1988. Equilibrium constant for the liquid-phase hydration of isobutylene over ion-exchange resin. Ind. Eng. Chem. Process Des. Dev. 25, 889e893. https://doi.org/10.1021/i200035a009. Douglas, J.M., 1998. Conceptual Design of Chemical Processes. McGraw-Hill, New York, USA. Garcia Jurado, M.B., Plesu, V., Bonet Ruiz, J., Bonet Ruiz, A.E., Tuluc, A., Llorens Llacuna, J., 2013. Simulation of a hybrid reactive extraction unit. biodiesel synthesis. Chem. Eng. Transactions. 35, 205e210. https://doi.org/10.3303/ CET1335034. Imam, R.A., Freund, H., Guit, R.P.M., Fellay, C., Meier, R.J., Sundmacher, K., 2013. Evaluation of different process concepts for the indirect hydration of cyclohexene to cyclohexanol. Org. Process Res. Dev. 17, 343e358. https://doi.org/10. 1021/op300276e. International Energy Agency, 2018. In: World Oil Demand. https://www.iea.org/ OILMARKETREPORT/OMRPUBLIC/. (Accessed 19 October 2018). Kiss, A.A., Jobson, M., 2018. Taking reactive distillation to the next level of process intensification. Chem. Eng. Trans. 69, 553e558. https://doi.org/10.3303/ CET1869093. Lee, H.Y., Lai, I.K., Huang, H.P., Chien, I.L., 2012. Design and control of thermally coupled reactive distillation for the production of isopropyl acetate. Ind. Eng. Chem. Res. 51, 11753e11763. https://doi.org/10.1021/ie300647h. Lucia, A., Amale, A., Taylor, R., 2006. Energy efficient hybrid separation processes. Ind. Eng. Chem. Res. 45, 8319e8328. https://doi.org/10.1021/ie060035t. Luyben, W.L., Yu, C.C., 2008. Reactive Distillation Design and Control. John Wiley & Sons, Inc., Hoboken, NJ. Malone, M.F., Huss, R.S., Doherty, M.F., 2003. Green chemical engineering aspects of reactive distillation. Environ. Sci. Technol. 37, 5325e5329. https://doi.org/10. 1021/es034467w. Norkobilov, A., Gorri, D., Ortiz, I., 2017. Comparative study of conventional, reactivedistillation and pervaporation integrated hybrid process for ethyl tert-butyl ether production. Chem. Eng. Process 122, 434e446. https://doi.org/10.1016/j. cep.2017.07.003. Novita, F.J., Lee, H.Y., Lee, M., 2015. Self-heat recuperative dividing wall column for enhancing the energy efficiency of the reactive distillation process in the formic acid production process. Chem. Eng. Process 97, 144e152. https://doi.org/10. 1016/j.cep.2015.09.007.

21

Novita, F.J., Lee, H.Y., Lee, M., 2017. Energy efficient design of ethyl levulinate reactive distillation process via thermally coupled with external heatintegrated arrangement. Ind. Eng. Chem. Res. 56, 7037e7048. https://doi.org/ 10.1021/acs.iecr.7b00667. Novita, F.J., Lee, H.Y., Lee, M., 2018a. Plantwide design for high-purity formic acid reactive distillation process with dividing wall column and external heat integration arrangements. Korean J. Chem. Eng. 35, 926e940. https://doi.org/10. 1007/s11814-017-0342-4. Novita, F.J., Lee, H.Y., Lee, M., 2018b. Reactive distillation with pervaporation hybrid configuration for enhanced ethyl levulinate production. Chem. Eng. Sci. 190, 297e311. https://doi.org/10.1016/j.ces.2018.06.024. ndez, J.G.S., Herna ndez, S., Aguirre, A.H., Yescas, R.M., 2012. Ojeda, M.V., Herna Optimization and controllability analysis of thermally coupled reactive distillation arrangements with minimum use of reboilers. Ind. Eng. Chem. Res. 51, 5856e5865. https://doi.org/10.1021/ie200929t. Sada, N.M., Kato, N.M., Mori, N.Y., Sano, N.M., 1981. Process for production of tertiary butyl alcohol. US Patent 4 (307), 257. Turton, R., Bailie, R.C., Whiting, W.B., Shaelwitz, J.A., 2009. Analysis, Synthesis, and Design of Chemical Processes, third ed. Prentice Hall, Upper Saddle River, NJ. U.S. Energy Information Administration, 2018. In: Energy Use in Industry. http:// www.eia.gov/energyexplained/index.cfm/data/index.cfm?page¼us_energy_ industry. (Accessed 19 October 2018). Velo, E., Puigjaner, L., Recasens, F., 1988. Inhibition by product in the liquid-phase hydration of isobutene to tert-butyl alcohol: kinetics and equilibrium studies. Ind. Eng. Chem. Res. 27, 2224e2231. https://doi.org/10.1021/ie00084a006. Wang, S.J., Wong, D.S.H., Yu, S.W., 2008. Design and control of transesterification reactive distillation with thermal coupling. Comput. Chem. Eng. 32, 3030. https://doi.org/10.1016/j.compchemeng.2008.04.001. Yu, H., Ye, Q., Xu, H., Zhang, H., Dai, X., 2015. Design and control of dividing-wall column for tert-butanol dehydration system via heterogeneous azeotropic distillation. Ind. Eng. Chem. Res. 54, 3384e3397. https://doi.org/10.1021/ ie504325g. Zhang, C.M., Adesina, A.A., Wainwright, M.S., 2003. Isobutene hydration over Amberlyst-15 in a slurry reactor. Chem. Eng. Process 42, 985e991. https://doi. org/10.1016/S0255-2701(02)00169-1.